A classic treatment on the removal of nitrogen oxides (NOx) from combustion generated gases using selective catalytic reduction (SCR) has been published by Bosch et. al. (1988), and more recently by Muzio et. al. (2002). Numerous patents have been awarded, with early work by Atsukawa (U.S. Pat. No. 4,302,431) typical of the initial status of the technology.
SCR is a relatively mature, well-developed environmental control technology for NOx emissions from fossil fuel fired power plants. The process has been extensively deployed internationally since the early 1970s, with approximately 180,000 MW of coal-fired boiler applications in the world at present. The first commercial SCR installations were deployed in Japan in the 1970s, initially on process heaters and refinery equipment, and subsequently on coal-fired power plants that fire by U.S. standards coal with low sulfur (0.5% or less) content. The mid-1980s witnessed the first commercial SCR applications in Europe, mostly confined to Germany, on coals typical of European usage. The origin of these coals from which European experience was generated was the Ruhr and Saar Valley in Germany, as well as Poland, South America, South Africa, Australia, and Venezuela. Most of these coals feature sulfur of 1-1.5% content. By the late 1980s, approximately 50,000 MW of SCR had been deployed on coal-fired units in Europe. In the U.S., approximately 100,000 MW of coal-fired capacity—over one third of the nation's coal-fired fleet—was operating by the end of the year 2004. The range of coals utilized are from extremely low sulfur western U.S. sources (e.g. Powder River Basin in Wyoming) to high sulfur sources such as Indiana, Pennsylvania, Ohio, Illinois and other midwestern U.S. sites.
The most problematic of these applications is for coal-fired power plants for the production of electricity, in particular those firing coals with sulfur content of 1.5% or higher. These applications are problematic due to the generation of two process byproducts. These are unreacted (residual) NH3 and sulfur trioxide (SO3), the former a consequence of incomplete reaction of NH3 with NOx, and the latter oxidized by the catalyst from SO2 in the flue gas. Both of these byproducts not only can comprise an environmental or nuisance hazard on their own, but can complicate the operation of balance-of-plant equipment. Most significantly, the oxidation of SO2 to SO3 has been problematic and responsible for the generation of visible plumes, receiving harsh public scrutiny (Akron Beacon Journal, 2001).
An innovative reactor design is described by Balling et. al. (U.S. Pat. No. 5,397,545) that uses a variable catalyst chemical composition to optimize the removal of NOx. Specifically, Balling et al. describe the concept of altering the composition of conventional SCR catalyst in the direction of flow, thus tailoring the chemical composition to the process conditions, to “accelerate” the reaction, in this case destruction of NOx. This art also introduces deploying a method of heat removal between layers, to arrest the escalation in temperature of the gas due to a reaction, with specific attention to the heat released by the exothermic reactions of CO oxidation. However, implementing this art recognizes only the potential to use a heat exchanger to correct for the complications introduced by upstream or preceding reactions. This art does not recognize the potential to use such a heat exchanger to establish a distinctly different process zone for a separate, downstream step that promotes optimal results of the entire process. The extraction of heat is particularly relevant in addressing design criteria for high sulfur coal which emphasizes the control of SO2 oxidation over maximizing the rate at which NO is removed from flue gas, an observation not recognized or exploited by Balling.
In addition to NOx, mercury emissions from coal-fired power plants have received significant attention in recent years. Mercury is introduced into boiler flue gas as a trace element naturally occurring in coal. Mercury will exist in boiler flue gas in either the elemental or oxidized state, with the most prevalent oxidized form either mercuric oxide (HgO) or mercuric chloride (HgCl2). The recent technical literature is replete with descriptions of investigations into control options for removing Hg from coal-fired utility flue gas (AWMA, 2001, 2003, and 2004).
Tests conducted by both the U.S. government and private utilities show that most of the Hg that exists in flue gas that is in a chemically oxidized state can be removed by conventional flue gas desulfurization (FGD) process equipment, installed for SO2 control (Chu, 2003). The consequences of this fact are significant—that technology installed for SO2 removal will also remove Hg—but only if the Hg is oxidized from the elemental form (Hge). Further, it is believed the FGD process chemistry can be manipulated in a manner to minimize the content of the sulfite ion in scrubbing liquid, in order to minimize the “re-emission” of captured Hg, and thus total Hg removal from the flue gas (Blythe, 2004).
The conventional deployment of an SCR process is relatively simple. FIG. 1 depicts the embodiment of an SCR process in a coal-fired power plant. Ammonia reagent (NH3) is injected 18 into the flue gas stream 20 produced by a boiler 22, where the flue gas is at temperatures ranging from 550-800° F. The mixture of ammonia reagent and NOx in the flue gas reacts within a catalytic reactor 24 in the presence of one or more layers of catalyst 26 that are specially prepared to reduce the NOx to molecular nitrogen and water.
The amount of NOx removed is directly proportional to the quantity of ammonia reagent injected, indicated by the normalized ratio of the moles of NH3 to the moles of NOx in the flue gas (the NH3/NO ratio). Equally important to the effective operation of the SCR process is avoiding a deleterious impact on the performance of the Ljungstrom regenerative air heater 30. This standard device, an integral component of essentially all coal-fired power plants, extracts residual heat from the flue gas before discharge through the stack. The performance of this air heater can be detrimentally impacted through residual NH3 and byproduct SO3, as will be described subsequently.
Of the many factors that affect SCR performance, the most important is the mixing of reagent NH3 with NOx in the flue gas. Several methods are pursued to maximize this mixing and achieve greatest SCR performance. FIG. 1 shows the location of devices installed in the ductwork, known as static mixers 28, leading into the reactor to improve the mixing of injected ammonia reagent with NOx, and assure that a relatively uniform distribution of flue gas velocity and composition reports to the reactor inlet. FIG. 1 depicts a popular location for the static mixers 28, which present a variety of surfaces of various geometrical shapes and orientation that impart mixing momentum to the flue gas.
FIG. 1 depicts the arrangement of catalyst within the reactor for the case of a three layer reactor design. For any given design, the total number of catalyst layers will vary with the type of fuel, desired NO removal, time between catalyst replacement, and other factors, and is usually between 2 and 5. Usually, all available layers that are provided in the catalytic reactor, such as depicted in FIG. 1, are not initially filled with catalyst. One layer is usually retained as a “spare” and filled with catalyst once the activity of the initial inventory begins to degrade, as a means to maximize the utilization of all catalyst.
For a given SCR process design and degree of NH3/NO mixing, the second most important practical limit to SCR performance is the volume and surface area of the catalyst within the reactor. This key design variable affects the residence time available for mass transfer and reaction. Even for good NH3/NO mixing and generous catalyst surface area, a small fraction of the injected reagent NH3 eludes contact with NO and migration to an active site, and thus does not react to produce the desired products of molecular nitrogen and water. The unreacted ammonia reagent is thereafter referred to as residual NH3. The NOx removal achievable with a given SCR design—be it 80%, 85%, or 90% or greater of inlet NOx values—is limited by the residual NH3 that is introduced into the flue gas. Combined with the production of SO3 from the catalyst, either or both of residual NH3 and byproduct SO3 can compromise the application of SCR.
In steam boilers for power generation, the SCR process is usually located as shown, between the boiler economizer section and the air heater. The temperature of flue gas at this location is well-suited for SCR application, as most boilers produce flue gas at a temperature between 500-800° F., adequate to provide the necessary catalyst activity for NOx removal. The maximum temperature at which the process can operate is determined by catalyst degradation, and a limit of 825° F. is usually observed.
The minimum flue gas temperature is determined by the flue gas sulfur trioxide (SO3) content, due to the reaction of SO3 with ammonia reagent in flue gas to form ammonium sulfates and ammonium bisulfates (ABS). These compounds can form at numerous locations following an SCR process, on surfaces downstream of the point of ammonia reagent injection, and as will be discussed throughout this disclosure, can be problematic for operation of both the SCR process and the power plant.
The potential to form ABS compounds on the catalyst and retard activity is of sufficient concern that boiler designers employ modifications to either the flue gas side or the steam side to insure that a minimum temperature is provided for, especially at lower loads where flue gas exiting the economizer is usually less than the values at full load. Recent contributions to the art of providing a minimum temperature for the SCR reactor at low load are taught by Cohen (U.S. Pat. No. 5,943,865), Ziegler (U.S. Pat. No. 5,775,266), and Wiechard (U.S. Pat. No. 5,555,849).
A summary of the flue gas temperatures for the onset of ABS deposits, as dependent on flue gas SO3 and NH3 content, is presented by “Ljungstrom Air Preheater Fouling Due To SCR Ammonia Slip” See attached reference for IDS, Counterman et. al. (1999) the subject matter of which is hereby incorporated by reference in its entirety. The highest temperatures noted (>500° F.) are those anticipated for a coal-fired power plant, due to the relatively high SO3 content generated from the sulfur dioxide (SO2) content by the boiler, and high NH3 concentrations corresponding to injected reagent. This temperature, defined by the relationship depicted on the third page of Counterman et al. (1999), establishes a minimum temperature “floor” for SCR operation, to prevent the deposition of ABS on the catalyst surface which will compromise NOx removal and damage the catalyst. The selection of minimum operating temperature must also recognize that condensation of ABS within the micropores of the catalyst must account for a different static pressure within the micropore, as induced by capillary action (Johnson, 2002, and Matsuda, 1982)
Process designers will modify the boiler to insure this minimum flue gas temperature is achieved at the reactor inlet, where the ABS deposition temperature is the highest that will be encountered. Subsequent to the first layer, the injected NH3 will have reacted and thus decreased in concentration, reducing potential for ABS deposition until the significantly lower temperatures of the air heater are encountered.
Only recently has the ability of SCR catalyst to oxidize elemental mercury in flue gas been recognized (Chu, 2002, and Laudel, 2003). Although first noted in the technical literature 12 years ago (Gutberlet, 1992), the significance was not recognized until the Information Collection Request (ICR) issued by the U.S. EPA in 1998-1999 was conducted to establish baseline mercury emissions from coal-fired power plants. Data from this program is in the public domain (EPA, 2000), and has been evaluated to provide insight as to how mercury can be controlled.
A number of investigations have attempted to isolate and measure the role of the SCR catalyst in increasing the oxidation of Hg in flue gas (Chu, 2002, Richardson, 2002, Chen, 2002, and Senior, 2004. Data suggests the propensity for SCR catalyst to oxidize elemental Hg is not always significant and consistent, and that many factors other than the SCR catalyst may be responsible. Among these factors is flue gas temperature, as Downs and co-workers at McDermott Technology Inc. have reported that operation of the catalyst at lower temperature may promote the oxidation of Hg from the elemental to the oxidized state (Downs, undated). Further, at least one investigator of the fundamental mechanisms of mercury oxidation has cited that altering the cooling rate between the economizer and air heater inlet can favor both thermodynamic and gas phase considerations and enhance the formation of mercuric chloride (HgCl2), believed to be the principal fate of oxidized mercury. (Chen, 2002). These sources suggest that a lower reactor temperature, extended residence time, and additional cooling steps are preferred for mercury oxidation.
The oxidized form of mercury, once produced, can be removed in a conventional flue gas desulfurization (FGD) process. Field test data shows that a large fraction of the oxidized mercury that enters an FGD process is removed, and that FGD process equipment preceded by an SCR process removes more mercury (Chu, 2002, Chu, 2003, Winberg, 2004, and Blythe, 2004). The specific amount of oxidized Hg removed depends on the composition of coal, and the design and operation of the FGD process, specifically the details of the sulfur chemistry. It has been noted the amount of Hg capture can vary significantly, ranging from less than 50% to as high as 90%. Recent investigations suggest the net mercury removal is the result of oxidized Hg first becoming solubilized and removed, but offset by a fraction of the same oxidized Hg reduced within the FGD process liquor back to the elemental state. This newly-generated elemental mercury is then “re-emitted” (Blythe, 2004). The details of the Hg removal chemistry are not clear, but Blythe and coworkers speculate that sulfite within the FGD liquor catalyzes the reduction of oxidized Hg to elemental. Accordingly, minimizing FGD sulfite content in solution provides the best environment for Hg collection. Manipulating FGD process chemistry to minimize sulfite content may offer an improved environment to minimize Hg “re-emission”, and improve net Hg capture.
In this regard, fundamental studies of FGD process chemistry conducted over 25 years ago suggests that NO2 and NO both can play roles in the oxidation sulfite within the FGD process. Specifically, Rosenberg (1980) showed in laboratory-scale tests that NO—by inhibiting the oxidation of sulfite to sulfate—acts to maximize the amount of sulfite in FGD solution. Rosenberg and coworkers also noted that NO2 acts to oxidize sulfite to sulfate, thereby minimizing the amount of sulfite in solution. This work was conducted to elucidate the role of NO and NO2 on sulfite oxidation, and considered NO to NO2 ratios that typify combustion products, specifically 12/1 to 19/1. Given the problem addressed at the time—controlling the oxidation of sulfite to minimize gypsum scaling—this work was used to conclude that FGD process chemistry benefited by a high ratio of NO to NO2, to prevent oxidation of sulfite to sulfate.
Several key design and operating factors affect the performance and operation of SCR process equipment, for which a typical relationship is summarized in FIG. 2. These factors, which as will be shown can limit the usefulness of the SCR process, are favorably affected by the inventive process. These factors are the generation of residual NH3 and byproduct SO3, which can significantly interfere with plant operations, as chronicled in early commercial U.S. SCR applications on higher sulfur coal (Akron-Beacon Journal, 2001).
FIG. 2 presents the key NOx control performance data, depicting the NOx removed and residual NH3 reagent introduced into the flue gas by the process. Specifically, FIG. 2 shows the typical approximate linear relationship between NOx removal (32, as indicated on the left y-axis) and NH3/NO ratio, shown on the x axis 34. The value of NOx removal is shown as an approximate linear relationship (36), and the amount of residual NH3 byproduct introduced into the flue gas on the right y axis (38).
Also shown in FIG. 2 are the residual NH3 values generated in exchange for this NOx removal. The value of residual NH3 will vary with remaining catalyst activity and thus lifetime, and is shown for three typical catalyst lifetimes: 10,000 hours 40, 16,000 hours 42, and 24,000 hours 44. The relationship depicted in FIG. 2 is for a hypothetical case of 200 ppm inlet NOx, but applies in principal to a wide range of process conditions.
FIG. 2 shows that almost any level of NOx removal can be attained—up to and approaching 95% of inlet values—depending on the level of unreacted NH3 in flue gas that can be tolerated. The key to controlling the residual NH3 is to maximize the mixing of injected reagent with flue gas, and the evolution of SCR technology is replete with attempts to maximize this mixing.
Most recently, a concept for a sophisticated injection system to insure balance between the injected reagent and combustion product gases has been applied for by Rogers et. al. (U.S. patent application 200330003029). As a further example of methods to improve this mixing, an injection grid has been described by Anderson et. al. (U.S. Pat. No. 5,988,115). An approach that is gaining widespread acceptance is the use of the static mixers that were shown in FIG. 1, which are important not only for stationery SCR applications as described by Henke (U.S. Pat. No. 4,737,345), but for mobile applications of SCR, as described by Hoffman et. al. (U.S. Pat. No. 6,553,755). Humsetal et. al. (U.S. Pat. No. 6,287,524) has devised methods to increase turbulence in the region of reagent injection to improve mixing. All of these approaches and others found in the literature teach improving mixing before entry into the reactor, but they do not address improving mixing once the flue gases have progressed beyond the first catalyst layer.
For most coal-fired boilers the concentration of unreacted NH3 that can be tolerated is limited to just 2-3 ppm, due to impact on balance-of-plant equipment such as the Ljungstrom air heater and other operating equipment. Accordingly, although in concept NH3/NO can be injected to achieve 95% NOx removal, practical limitations that restrict residual NH3 to 2-3 ppm constrain NOx reduction to 85-93%. Significantly, FIG. 2 also shows that the residual NH3 generated depends on the catalyst lifetime, as residual NH3 increases as catalyst lifetime increases from 10,000 hours 40, to 16,000 hours 42, to 22,000 hours 44.
The actual NH3/NO ratio entering any catalyst layer is not a single value, but a distribution of values, each of which can significantly deviate from the average. SCR equipment is designed to meet a process specification, which defines among other factors the variance in flue gas concentration of NH3/NO and other key process factors, such as NOx, flue gas temperature and the distribution of velocity at the inlet of the reactor. This variance is usually defined in terms of the standard deviation of a set of values from a mean.
For any given process specification, catalyst of sufficient volume and composition is provided to remove NO and control unreacted NH3 reagent from flue gas, at a specified temperature. As an example, an SCR process specification may require a certain performance target from flue gas with an average velocity distribution of 15%, and NH3/NO distribution of 6%, and temperature variance of 30° F.
In general, the most important of these SCR variables is the distribution of NH3/NO. The significance of high standard deviation in NH3/NO is that it creates local zones both less and greater than the NH3/NO ratio required to react. Specifically, a portion of these local zones in the reactor will experience extremely low NH3/NO ratio, well below the mean value, and other portions of the reactor will experience NH3/NO ratio above the mean. Both contribute to compromised performance. The lower value NH3/NO zones do not maximize the use of the catalyst, and the higher value NH3/NO zones provide excess reagent over the quantity of NO injected. This NH3 reagent in excess of the NO in the flue gas cannot react, and will generate residual NH3 and thus limit the operation of the entire process.
The influence of changes in the standard deviation of NH3/NO is shown in FIG. 3, showing the relationship between NOx removal achieved 50 and residual NH3 52, for the specified case of 200 ppm inlet NOx, and flue gas temperature of 700° F., inlet reactor NH3/NO ratio of 0.91. The data in FIG. 3 shows that if 5% standard deviation of NH3/NO can be achieved 54—reflecting good mixing and allowing operation at a NH3/NO ratio of 0.91—NO removal of about 90% can be achieved for a residual NH3 limit of 2 ppm. However, if the standard deviation of NH3/NO is only 10% 56—then NOx removal must be limited to about 86% (by lowering the NH3/NO ratio to 0.85) to maintain a residual NH3 limit of 2 ppm. FIG. 3 demonstrates why SCR process designers utilize static mixers and other means to minimize NH3/NO standard deviation in selecting process design.
FIGS. 2 and 3 depict process inlet and outlet data, but do not describe process conditions across each layer that limit SCR performance. A significant characteristic of the SCR process is the progression of NOx removal and consumption of NH3 through the catalytic reactor. FIG. 4 presents data describing how key variables change across each layer of a three layer reactor, calculated for an example case similar to that of FIG. 3. FIG. 4 presents example calculations for an inlet NOx concentration of 200 ppm, operating at an NH3/NO ratio of 0.91, and achieving a 90% NOx removal. The key variables shown are NO removal, the NH3/NO ratio, and the standard deviation of NH3/NO ratio, at the inlet of a conventional design reactor. FIG. 4 also shows how these key variables change across each layer. FIG. 4 shows that 68% of the NOx is removed across the first layer, 19% across the second layer, and only 3% across the last layer. In fact, due to the relatively small amount of NOx removed across the last layer, its purpose as limited by conventional SCR design is as much to reduce unreacted NH3 to negligible levels, as contribute to NO removal. Of note is that the NH3/NO ratio calculated for each of the three layers decreases through the reactor. The first layer is exposed to an NH3/NO ratio equivalent to the inlet of the process reactor—in this case 0.91.
However, the simultaneous consumption of ammonia and NOx lowers the NH3/NO ratio to 0.72 at the exit of the first layer, which corresponds to the inlet NH3/NO of the second layer. This same pattern continues, with NH3/NO leaving the second layer and thus entering the third layer as 0.31. Consequently, the last catalyst layer, although contributing equally to the capital and operating cost of the process as much as the first layer, contributes relatively little NOx removal.
Significantly, as the reaction progresses through each layer in the catalytic reactor, small deviations in NH3/NO in the first layer translate into large deviations in NH3/NO in subsequent layers. FIG. 4 shows the deviation in NH3 that is observed at the reactor inlet, based on a 5% standard deviation at the reactor inlet. For these process conditions, the magnitude of the 5% unmixedness equates to about a 9 ppm excess or deficit of reagent at any point. This deviation from mean values is unchanged at 9 ppm as the reaction proceeds, but the fraction this 9 ppm of NH3 represents of remaining NO increases. Accordingly, for the example case, the standard deviation of the NH3/NO ratio entering the second and third layer increases, respectively, to 14 and 35%.
Thus, the practical NOx removal is limited by the ability to contact essentially all injected ammonia reagent with NOx, at an active catalyst site.
In general, most SCR catalysts oxidize from 0.5 to 3% of the SO2 contained in the flue gas to SO3. FIG. 4 included an estimate of the increase in SO3 content across each layer for the example case. This extent of oxidation is dependent on the features of the catalyst, and the volume of catalyst applied.
Conventional practice allows for both catalyst and process design to minimize the conversion of SO2 to SO3. Some catalyst suppliers substitute for vanadium other active materials such as molybdenum that can provide NO removal (although less than compared to vanadium) but minimize SO2 oxidation. In this way, SO2 conversion can be reduced from 2% or greater to less than 1%. However, a larger catalyst volume to provide the same NO removal is usually required, to compensate for the lower activity with respect to NO removal.
The second method to control SO2 conversion is lowering reactor operating temperature. FIG. 5 presents a typical relationship describing SO2 oxidation rates for a commercial catalyst as a function of flue gas temperature. As shown, the rate of SO2 conversion 60 is strongly dependent on flue gas temperature 62, with a reduction by 50° F. lowering the conversion rate by almost a factor of two 64 to 66. This relationship is exploited by process designers in minimizing SO2 conversion for a particular application. Several commercial process designs have been noted where the boiler is modified so that the flue gas in the economizer exit section—where the SCR process is installed—is lowered by 50 to 75° F., thus mitigating SO2 conversion.
The detailed physics of SO2 to SO3 oxidation suggests that, similar to the case of NO removal and ammonia reagent consumption, process conditions across each layer can vary significantly. Knowledge of this layer-by-layer variation can be exploited to provide a low SO2 conversion reactor compared to conventional reactor design. Specifically, statistical thermodynamics dictates that all reactants in flue gas, such as NO, NH3, SO2, and O2—compete with each other (as well as background species of CO2, H2O, etc.) for access to active catalyst sites. Accordingly, the depletion of NO and NH3 improves access of SO2 to an active site, increasing SO2 oxidation. This phenomena has been observed in the laboratory and at pilot scale—that SO2 conversion is higher where the concentration of NO and NH3 are relatively low. At laboratory scale, tests conducted by Svachula (1993) measured the influence on SO2 oxidation of a large number of process variables. FIG. 6 shows the experimental relationship between SO2 oxidation on the left-y axis 70 and NH3/NO ratio on the x-axis 72. Data in this figure shows that SO2 oxidation, at low values of NH3/NO ratio such as 0.2 74, is about a factor of two greater than SO2 oxidation when NH3/NO ratio is at typical operating levels of 0.80 76.
These experiments suggest that as flue gas passes through the reactor, each catalyst layer contributes an increasing conversion of SO2 to SO3 to the entire process. The last layer, where the concentration of NH3 and NO is the lowest, contributes the most to the overall reactor oxidation of SO2.
FIG. 4, in addition to summarizing NO and NH3 concentration within a reactor, also reported the increase in SO3 content across each layer, based on trends identified by Svachula. Specifically, FIG. 4 assumed the SO2 oxidation is 0.3%, 0.5%, and 0.7% for the first, second, and third layer, respectively. The amount of SO3 across each catalyst layer is shown, based on the amount created across each layer for a flue gas SO2 content of 2,000 ppm, added to the inlet values produced by the boiler. Also shown is the temperature of deposition of ABS for each layer, as dependent on the NH3 and SO3 in the flue gas. This data shows that the ABS deposition temperature decreases with each layer. The highest ABS deposition temperature is for the first layer, and this establishes the minimum reactor operating temperature, for the conventional SCR design.
The generation of residual NH3 and byproduct SO3 from SCR can adversely impact operation of the entire power station.
Of note is that introduction of residual NH3 by itself into the flue gas in quantities typical of SCR does not necessarily cause harm. There are possible environmental impacts, but these appear to be at concentrations well above those typical for flue gas at the SCR process exit. However, the secondary impacts on power plant and balance-of-plant equipment are of considerable concern. For residual NH3 these are contamination of ash, and in conjunction with SO3 air heater plugging from ammonium sulfates and bisulfates and materials corrosion.
Residual NH3 will be absorbed onto the fly ash, and can compromise the sale of fly ash for construction supplement, or acceptable disposal. Early experience (1980-1985) in Japan suggested that limiting unreacted NH3 in flue gas to 5 ppm minimized absorption of ammonia by fly ash and avoided these problems. However, this threshold—established for the types of coals fired in Japan—was not adequate for the conditions of application in Europe. German experience (1986 and after), consistent with early lessons from the first operating U.S. applications (1991 and on), showed that depending on the particular use for fly ash, the flue gas residual NH3 should not exceed 2 ppm. This lower flue gas content maintained ammonia concentration in the ash generally below 100 ppm. Accordingly, the generally accepted design threshold for flue gas residual NH3 was modified in the late-1980s to be 2-3 ppm. This design limit has been applied to the majority of U.S. installations, and maintaining residual NH3 below this limit (along with avoiding air heater plugging, as discussed in the next section) establishes when either NO removal must be compromised, or catalyst added or exchanged.
The immediate and significant consequence of flue gas SO3—as formed either inherently in coal-fired systems or augmented by the presence of SCR—is the production of ammonium sulfates and bisulfates from residual NH3 and byproduct SO3. As described previously, ammonium sulfates and bisulfates can form in the reactor, and specifically are most prone to form on the first layer where the injected NH3 concentration is high. The flue gas temperatures where the deposition can occur have been previously described in FIG. 4. Once the injected NH3 reacts and the concentration is reduced, ammonium sulfates and bisulfates are less likely to form at the temperatures typical of an SCR reactor (640-700 F), but could form as the flue gas cools to 400-500 F as it passes through the Ljungstrom heat exchanger.
FIG. 7 represents a conventional Ljungstrom air heater for use in power boilers. The purpose of this device is to recover the last amount of usable heat prior to entry of the flue gas to the environmental control system. Flue gas enters a heat exchanger shell 80, that is split in two sections, segregating flue gas exiting the boiler from combustion air entering the boiler. The gases pass through three stages of heat exchange elements that are aligned in the direction of flow. These heat exchange elements are assembled as “baskets” that can be easily removed and replaced, and are essentially a series of plates spaced to allow flue gas flow with minimal pressure drop but the necessary heat transfer characteristics. The three stages of the heat transfer baskets are: the “hot” 82, “intermediate” 84, and “cold” 86 sections. These baskets rotate at approximately 1 rpm between the flue gas and combustion air sections. Hot flue gases from the boiler give up heat to the baskets, which retain this heat and rotate into the combustion air, subsequently increasing the temperature of the combustion air on the way to the boiler.
Essentially all commercial boilers for power generation utilize such a device, and in excess of 90% employ the Ljungstrom-type design. The major design variant is whether the axis of rotation of the heat exchange material is horizontally or vertically configured. The device usually lowers the flue gas temperature from approximately 575-725° F. (the specific value depending on boiler design and coal composition) to 350-275° F., transferring the heat to the incoming combustion air, and improving the combustion process. To maintain reliable and effective performance, the heat exchange materials or “baskets” should remain clean and unobstructed, otherwise the flue gas pressure resistance will increase beyond the design capability, restricting the maximum flue gas flow rate and thus boiler power output.
FIG. 8 shows in more detail the three elements of the heat transfer surfaces: the “hot”, “intermediate”, and “cold” sections. These surfaces are generally kept clean through the use of “sootblowing”, in which a high pressure (100 psig) pulse of cleaning media (usually steam from the boiler) is injected across the face of the surfaces. The sootblowing lances can be located at either the hot-end 96 and thus inject cleaning media in the direction of flue gas flow, or the cold-end 98 and thus inject cleaning media against the direction of flue gas flow. The steam or other cleaning media is injected via nozzles that are swept across the inlet of the baskets, covering the entire cross-section in 1-3 minutes. The frequency with which this cleaning process is repeated varies widely, from several times per day to once per month, depending on the type of coal fired, and boiler design.
Each of these heat exchange sectors is characterized by a different flue gas temperature range. The hot section generally lowers the flue gas from the reactor inlet values of 575-725 to 500-575° F. At these relatively high temperatures there is little opportunity for formation of ABS compounds, and subsequent deposition on heat exchange surfaces. Accordingly, the materials chosen for fabrication of the hot-end baskets 90 can be lower cost conventional steels, and the spacing of plates can be minimized, to maximize heat transfer without concern for plugging. The intermediate baskets 92 lower flue gas temperature from 500-575 to 400-450° F. In this section, the opportunity for deposition and plugging induced by lower temperatures improves, thus materials for some applications should be constructed of corrosion-resistant alloys. Also, the spacing between plates may be required to be increased, to minimize the opportunity for plugging passageways. The cold-end baskets 94, as the last element, lowers flue gas from 400-450 F to the exit of 275-350 F, depending on the fuel and design of the boiler. It is this cold end section that experiences the most aggressive conditions with respect to plugging by deposition of ABS, as well as SO3 deposition to form condensed sulfuric acid. It should be noted the cited temperature ranges are approximate, and individual applications can vary depending on boiler performance and coal composition.
An additional operating characteristic of the Ljungstrom-type air heater is the cyclic temperature variation that the heat exchange materials or baskets experience. This distribution in temperature is important as it determines the condensation of flue gas SO3 into sulfuric acid. Specifically, FIG. 16 illustrates the calculated surface temperature of the heat exchange baskets, as a function of the depth within the air heater. In FIG. 16, the zero depth corresponds to the inlet or hot side of the air heater, and increasing depth as measured in inches represents the flow direction toward the cold-end or exit of the air heater, which is at 52 inches for the example case cited. FIG. 16 illustrates that the surface temperature of the heat exchange materials varies between and a minimum and maximum. This temperature variation is due to the cyclic movement of the heat exchange baskets from the relatively higher temperature flue gas (to be cooled) into lower temperature combustion air (to absorb heat). As shown in FIG. 16, at any given distance into the air heater, a maximum surface temperature 170 is observed as the basket emerges from the relatively high temperature flue gas into the cooler combustion air, and this surface eventually gives up heat to the cooler combustion air, and experiences a minimum temperature 174, prior to re-entering the flue gas. FIG. 16 also reports an average surface or metal temperature 172, determined from the maximum and minimum surface or metal temperatures shown. FIG. 16 shows the minimum cold-end metal temperature is lower than the average cold-end metal temperature, by anywhere from 72° F. at the air heater inlet (at zero inches depth) to 32° F. at the air heater outlet (at 52 inches depth). The power plant designer is concerned with both the minimum and the average metal temperatures, at the maximum depth into the air heater, corresponding the exit plane, known as the cold-end. In fact, the power plant designer, in seeking to maximize power plant thermal efficiency, specifies a cold-end average metal temperature that is the lowest possible without incurring unacceptable SO3-derived corrosion. Design decisions must address not only the cold-end average metal temperature but also the minimum metal temperature, as the latter provides the best opportunity for condensation of sulfuric acid, the prime actor for corrosion. Accordingly, a method to enable a Ljungstrom-type air heater to extract the same or more heat from flue gas while minimizing the variance in minimum cold-end metal temperature from the average cold-end metal temperature would represent an improvement to the state of art of air heater design. Designers could exploit this in either of two ways: the first by retaining the same average cold-end metal temperature with a less extreme minimum temperature, thereby reducing SO3-derived corrosion; the second by retaining the same minimum cold-end metal temperature, but exploiting a lower average metal temperature thereby increasing thermal efficiency. Consequently, in the first case, SO3-derived corrosion would be reduced for the same boiler thermal efficiency; in the second case the boiler thermal efficiency would be increased for the same level of SO3-derived corrosion.
Research during the early period of SCR evolution showed that residual NH3 from SCR will combine with SO3 in the flue gas to form ammonium sulfates and bisulfates at the interface of the intermediate and cold-end sections, as depicted in FIG. 8. The occurrence of these deposits at this interface appears to be a consequence of both the temperature and physical conditions of ammonium sulfates and bisulfates formation, and the effectiveness of the cleansing from sootblowing activities. Specifically, it is believed that deposits occur at the interface of the intermediate and cold ends sections because the “void” 95 between the two sections allows the high energy cleansing jet of steam or other cleaning media to expand, diffusing the cleaning momentum into the void, thus compromising the ability to maintain the surfaces clean. A solution tried with some success is integrating the intermediate and cold sections into a single combined intermediate/cold section (Bondurant, 1999). The advantage of the combined cold/intermediate section is that eliminating the void between the intermediate and cold-end sections prevents the soot blower medium from diffusing and compromising cleaning momentum. The details of this mechanism have never been proven, but experience with one piece, intermediate/cold-end baskets on SCR-equipped units demonstrates these surfaces can help minimize, but not completely eliminate, accumulation of ammonium sulfates and bisulfates.
The reason why air heater deposition problems are not completely eliminated but persist with units equipped with a one piece intermediate/cold-end section is that the sootblowing media must still traverse the void between the hot-end and the one piece intermediate/cold sections. This remaining void can still dissipate the sootblowing media momentum, compromising cleaning ability. Ideally, a one piece heat exchange element combining the hot, intermediate, and cold-end sections could minimize or even eliminate the dissipation of sootblowing momentum. However, manufacturing limitations prevent such a one piece element from being constructed, at least for a cost that is competitive with the present approach. A number of technical and cost barriers prevent configuring all three sections (hot/intermediate/cold) into a one piece element. Recent contributions to the art of air heater design by Fierle et. al. (U.S. Pat. No. 6,260,606) are typical of attempts to mitigate the accumulation of ABS in the cold end section, as are the techniques described by Bondurant (1999).
As noted previously, the Ljungstrom-type air heater is the predominant design type of air heater used in power boiler applications, and the background information presented to date has rightly focused on this most popular concept. Other types of air heater designs are used, among them the Rothemuhle design, which is popular particularly in applications outside the U.S. The Rothemuhle design differs from the Ljungstrom concept in that the heat absorber plates, rather than rotating, are stationary. It is the rotational action of a flue hood that diverts flue gas from the combustion products and the air entering the boiler into different sections for heat exchange. The challenges imposed by SCR on air heater operation for the Rothemuhle design are conceptually the same as the Ljungstrom design, although the details are different.